Ind. Eng. Chem. Res. 1994,33, 2421-2425
2421
Modeling and Simulation of a Three-phase Catalytic Membrane Reactor for Nitrobenzene Hydrogenation Miguel Torres,+Josh Sanchez,t Jean A. Dalmon,t Bohumil Bernauer,s and Joseph Lieto'J Laboratoire d'Automatique et de Gknie des Proc&d&sLAGEP URA CNRS D 1328, CPE-Lyon et Universitk Claude Bernard Lyon I, Bat 721,43, Bd du 11 novembre 1918, 69622 Villeurbanne Cedex, France, Znstitut de Recherches sur la Catalyse, CNRS, 2, Av. A. Einstein, 69622 Villeurbanne Cedex, France, and Department of Inorganic Technology, Institute of Chemical Technology, Technica 5, CS-166 28 Prague, Czech Republic
The modeling and simulation of a catalytic membrane reactor have been studied for nitrobenzene hydrogenation to aniline in a three-phase system. A complete study of mass transfer mechanisms in different reactor configurations is proposed. The simulation of concentration profiles through the reactor was obtained by solving mass transfer balances corresponding t o different control volumes and including a kinetic rate obtained from independent experiments. The proposed model is in good agreement with experimental results and predicts that catalytic membrane reactors can be efficient for low-temperature three-phase reactions.
Introduction
Table 1. Catalyst Characteristics and Operating Conditions for the Kinetic Determination
Many important fine chemicals are synthesised by threephase catalytic reactions. In general, these reactions are carried out in different types of reactors such as stirred tank, fixed bed, or fluidized bed. In all cases, one of the controlling parameters is the mass transfer between gas and liquid reactants and/or the solid catalyst. For instance, it is well-known that, for hydrogenation reactions, the gas solubility in the liquid phase is in general very poor. Furthermore, the catalyst particles are surrounded by a liquid film and the mass transfer resistance is very important (Mills and Dudukovic, 1984). In order to decrease the mass transfer problems, a number of investigations have concentrated on the reactor configuration. Authors developed the idea to separate gas and liquid reactants by a porous wall (catalyst plus support). Many configurations have been studied: cross flow reactors (de Vos and Hamrin, 1982), monolithic reactors (Hatziantoniou et al., 1984), and membrane reactors (Cini and Harold, 1991). The most important conclusion of these previous works is that the new configurations are more efficient than a classical trickle bed reactor. Reactor modeling has been established, but only a limited number of parameters were studied experimentally and the simulations have not been thoroughly validated. For example, Akyurtlu et al. (1988) developed a semianalytical modeling coupling analytical results with a numericalsolution. In this work, we propose a complete modeling of a catalytic membrane reactor for a three-phase model reaction, the nitrobenzene hydrogenation. Pi.
The three-phase reaction mass transfer modeling we propose is completely numerical and coupled with experimental parameters such as membrane characteristics and reaction kinetics. The numerical simulation is compared with experimental results.
* To whom correspondence should be addressed. + CPE-Lyon et Universitb Claude Bernard Lyon I.
* Institut de Recherches sur la Catalyse. 8
Institute of Chemical Technology.
particle diam Ptly-Alzos (WIW) Pt particle d i m Pt dispersion mean pore diam specific surface pressure temperature total volume amount of Pt nitrobenzene concn shaft rotation
(20-200) x 10" m 7.0% (1-2) x 1O-em 80% 6X10"m
250 m2.g-l (1-5) X 106 Pa 303-333 K 80 X 1o-B ma (0.7-5)X l@ kg (1.33-6.11) X l(r mol/ ( 1 ~m0i/m3) 0 17 Hz
Experimental Section Reaction Kinetics. The reaction kinetics was studied in a stirred tank reactor (Burton Corblin, 100mL stainless steel tank equipped with a self-inducing turbine allowing an efficient gas recirculation in the liquid phase). For this purpose, we used a catalyst (PtlyAl~Oa)prepared by ionic exchange with HzPtCb. Before each experiment, the catalyst was activated "in situ" with hydrogen at 673 K. The experimental conditions and catalyst characteristics are reported in Table 1. We studied the influence of hydrogen pressure, nitrobenzene concentration, and temperature. We obtained the following kinetic expression: ra = 4.0 X 10' exp( -5* 6 X lo3)C*HMR (1) T (K) Details of the experimental procedure are given by Torres (1993). Catalytic Membrane Reactor. The experimental setup for the catalytic membrane reactor is shown in Figure 1. It consists essentially of a stirred tank and a tubular membrane reactor. The liquid phase (ethanol solution of nitrobenzene) is recirculated, and sampling is performed for chromatographicanalysis purpose (Intersmat 120FID chromatograph, equipped with an Altech C-500column with nitrogen as gas carrier). The membrane reactor is tubular with one composite membrane of the same geometry inside. This configuration defines two different reactor compartments: the most internal is the inner compartment and the space placed between the membrane support and the reactor wall is called the outer compart-
0888-5885/94/2633-2421$04.50/00 1994 American Chemical Society
2422 Ind. Eng. Chem. Res., Vol. 33, No. 10, 1994
rn
s u m
liquid phase
at
M m h c
D,
Figure 3. Schematic diagram of the membrane reactor and cwrdinates: (a) lateral view: (h) front view. Figure I. Schematic diagram of the experimental setup.
Configuration 1
gas
I
Configuration 2
I liquid
Figure 2. Configurations1 and 2.
ment. Liquid or gas can be fed in the reactor through the two compartments defining two configurations: liquid in the outer compartment and gas in the inner compartment correspond to configuration 1. The opposite situation corresponds to configuration 2 (Figure 2). The temperature in the stirred tank and membrane reactor is maintained constant by recirculating water in the corresponding jackets. The alumina membrane, supplied by SCT (US Filters), is composed by two different layers. The most internal in y-alumina (mean pore size of 5 X l W m and is 2.5 X 1W m thick). A macroporous a-alumina tube (1.6 X 1lV m thick, 7.0 X 10-3 m inner diameter, and 0.25 or 0.4 m long) is the support. The tube is glass sealed a t the ends. Platinum was used as catalyst and deposited in the membrane by a classical ionic exchange (HzPtCk). A reduction step with hydrogen at 673 K was performed. The membrane was characterized by different techniques (SEM, atomic absorption, Nz absorption) (Uzio et al., 1993). The most important characteristics are given in Table 2. The experiments are performed with the following procedure: the stirred tank is filled with an ethanol solution of nitrobenzene and recirculated continuously between thereactorchamber and thetank. Hydrogenisintrcduced in the other compartment. Some experiments were performed with total saturation of the liquid phase by hydrogen. For these experiments, the gas was introduced simultaneously in the membrane reactor and the stirred tank. Differential pressure between the two reactor compartments is maintained at zero, and samples are picked up and analyzed continuously. With these experimental conditions, no liquid leakage was observed for both configurations.
Table 2. Membrane Charactadstics inner diam external diam length mean pore diam of a-AI*Os support support thicknew mean pore diam of r-AlzOa layer .r-AlzO, thickneaa porosity (Y-AIzOS) porosity (a-AlzOs) tortuosity (.r-AIzOd tortuosity (a-AbOa) Pt contentI.r-Al20s(w/w) Pt particle d i m specific surface (.r-Al?Os)
0.1 x 1Cr2 m 1.1 x
1Vm
(26-40) X
lCr2 m
1Ox104m
1.1 x 109 m 5X1Crsm
(2-3) X 104 m 48% 30%
2.0 1.5
1% (1-3) X 104 m 150 m2.g'
The first studied variable was the influence of the hydrogen saturation of the liquid phase. We observed that a t 293 K and for configuration 2 (liquid in the inner compartment) the conversion obtained when saturation is performed is 30% higher than that obtained with a nonsaturated liquid phase. This result can be explained by the fact that a noticeable part of hydrogen consumed by the reaction comes from the liquid phase. The support is completely filled with liquid, and the gas has to diffuse through the total alumina tube thickness before reaching the catalytic active sites. For configuration 1(gas in the inner compartment), we obtained the same conversion values whatever the liquid phase hydrogen content is. The conversion is the same as the one obtained with the liquidsaturation inconfiguration 2. This result suggests that for this configuration (gas in the inner compartment) the hydrogen consumed during the reaction comes mainly from the gas phase. However, numerical simulation, showed that, for configuration 1 andsaturated by liquid, hydrogencomes from bothphases. Therefore a hydrogen concentration gradient through the membrane and the support is created.
Model Development In the previous studies of Akyurtlu et al. (1988) and Harold et al. (1989) the modeling of catalytic membrane reactors for three-phase reactions have been developed. The present modeling is based on a different system. We consider here a two-layer membrane: a mesoporous catalytic membrane (2.5 X 1W m thick) supported on a macroporous catalytically inert cylindrical wall (1.6 X 1lV mthick). Thechemicalreactionsoccur onlyinthecatalytic layer with simultaneous mass transfer of reactants in the support and the liquid channel. Furthermore outputmean concentrations of reactants are calculated and the dynamic behavior of the closed-loop reaction system is completely modeled. The cylindrical geometry of the reactor defines the coordinates shown in Figure 3.
Ind. Eng. Chem. Res., Vol. 33, No. 10,1994 2423 Table 3. Dimensionless Boundary Conditions for Configurations 1 and 2
C”=1
configuration 1
dC,’ -= dtm
0
configuration 2
The following assumptions and simplifications were taken into account to develop the differential equations system for the analysis (Torres, 1993): isothermal conditions; fully developed laminar flow in the liquid phase; membrane and support pores filled with liquid by capillarity; constant gas and liquid physicochemical characteristics; steady state regime; consideration of the stirred tank is considered as an ideal stirred vessel. The steady state mass balances for the system are given as follows:
following definitions:
r1 r3 6, = r3 - r2, Y~ = - = ‘m ($HZ
= KC,* -a m2 DHm
in the membrane (rl < r < r2): 1 d dcim D. -- r- R i = O (i=H,B) Imrdr( dr
)
in the support (rz< r
‘3
($B2
(2)
Kc,, 2 =DBmcH*
We can write now the dimensionless form of the equations:
< r3): (3)
in the membrane:
The conservation equations in the liquid channel are
where in the support: for configuration 1 and
For the liquid channel we can write for configuration 2. We take into account the boundary conditions (BC) at the gas-liquid interface as with dCB --0 dr Between the support and the membrane, the BC can be written as
73
= 0 for configuration 2, and
Furthermore, the BC for channel equations (4) can be expressed as dC, aciL D, -= Di - (configuration 1) dr ar
(7)
aCiL dCim Dim-- Di - (configuration 2) ar dr
(8)
Writing the mass balance equations for the different control volumes in dimensionless form requires the
73
= r3/63
(19)
for configuration 1. Boundary conditions for both configurations are summarized in Table 3. For the liquid channel the dimensionless output concentrations (2= 1)were computed as mean mixing-cup values:
2424 Ind. Eng. Chem. Res., Vol. 33, No. 10, 1994 Table 4. Parameters Used in Numerical Simulations internal radius 0.35 X m 0.325 X 10-2 m membrane radius 0.55 X m support radius 0.65 X m reactor radius 6.11 X mol/(mol/m3) nitrobenzene concn Ha diffusivity 2.7 X 10-9 m2/s effective diffusion coeff (Ha) 5.4 x 10-10 m2/s nitrobenzene diffusivity (1-9)x 10-10 m2/s 1.8 X 10-lo m2/s effective diffusion coeff (nitrobenzene) 3 stoichiometric coeff (Hz/ nitrobenzene) membrane porosity 48% membrane tortuousity 2.0 support porosity 30% support tortuousity 1.5 no. of collocation points 8
1
0,s
membrane
eupport
u
%01 ' ' ' ' ' ' ' 06'
'
'
'
8
8
1
8
1
1
16
10
20
Radial Position
Figure 4. Simulated concentration profiles for hydrogen at different Thiele moduli.
where G ( ~ Lis) the dimensionless laminar velocity distribution in the channel. These values of Cil were used in the mass balance of the ideal stirred tank.
Initial conditions (CiL, ( 0 , t ~ ) for ) the channel equations are related with the stirred tank output concentrations, CL: Z = 0,C~L(O,~L)= CiL. The hydrogen concentration in the stirred tank is constant (CHLO = 1) when the liquid is continuously saturated by H2 or = C H L ( ~in) the opposite case (the stirred tank is closed). The initial nitrobenzene concentration (CBL(O,~L)) is determined from the stirred tank mass balance. Numerical Simulations. The governing equations (11-14)were solved numerically by the orthogonal collocation method on two spatial dimensions (Finlayson, 1980). The solution was obtained simultaneously in the axial and radial coordinates. We could simulate with this method the hydrogen and nitrobenzene concentrations in the liquid phase, in the membrane and the support, for the two configurations. Diffusivity values for hydrogen in the liquid phase were estimated by the Wilke-Chang equation (Sherwoodet al., 1975)whereas the experimental kinetics were used. A list of the parameters used in the numerical simulation is given in Table 4. The results of the numerical simulation are given in Figures 4-6. Figure 4 illustrates the effect of the reaction regime associated with the values of the hydrogen Thiele modulus ( 9 =~1 and ~ &.i2 = 10) on the concentration profiles in the reactor functioning in configuration 1 and for the liquid saturated with hydrogen. When the diffusional regime is prevailing, the hydrogen concentration decay is very important because the mass transfer (diffusion of hydrogen in the liquid phase) is very slow: almost all molecules have a chance to react on active sites. Furthermore, the simulation shows that the diffusion through the support is not negligible when the pores are filled with the liquid. Figure 5 shows the hydrogen and nitrobenzene concentration profiles when the measured reaction kinetics is used for configuration 1. We can observe that hydrogen concentration decreases less than nitrobenzene concentration. If we compare this result with the previous one, we can conclude that the reaction occurs in an intermediate regime associating mass transfer with chemical reaction. In all cases, the diffusion of hydrogen and nitrobenzene
liquid
-,-
0
0.2 0.4 0.6 0.8 1.0 1.2 1.4
I
1.6 1.8 2.0 2.2 2.4 2.6 2.8 3.0
Radial Position
Figure 5. Calculated hydrogen and nitrobenzene concentration profiles (configuration 1 and at T = 313 K).
Membrane
Liquid
1
Support
1
Figure 6. Tridimensional representation of hydrogen concentration profiles (configuration 2 and at T = 313 K).
through the support is a very important parameter for the two following reasons. 1. The resistance to the diffusion in the liquid phase is not negligible. 2. The support is very thick when compared with the membrane itself (Table 2). A tridimensional representation of hydrogen concentration profiles is shown in Figure 6 for configuration 2. We can observe that the hydrogen concentration profile is more or less constant through the reactor length. This behavior can be explained by the fact that the conversion is very low for each space time through the reactor (T = 7-15 8). Validation. To validate the model, we assumed that the membrane reactor was a plug flow reactor and the tank was an ideal stirred vessel. This assumption can be used because the space time is very short when compared with the reaction time.
Ind. Eng. Chem. Res., Vol. 33, No. 10, 1994 2425
.-c
0.7 ~0 1.0
2.0
3.0
4.0
6.0
Time (h)
Figure 7. Comparison between experimental and simulated results at 313 K and for configurations 1 and 2 (-, experimental).
simulated; e,
Figure 7 shows the dimensionless concentration values obtained by simulation and experimentally as a function of the reaction time a t 313 K. The first important result is that the membrane does not present any loss of activity during 60 h of reaction, giving aniline as the only hydrogenation product. The other important conclusion is that the conversion values obtained by calculation are in good agreement with experimental results. At T = 313 K, the two configurations 1 and 2 produce the same conversion because, in both cases, the support and membrane pores are filled with liquid. Therefore, hydrogen and nitrobenzene diffusions prevail.
Conclusion We studied the nitrobenzene hydrogenation in a new type of three-phase catalytic membrane tubular reactor. We demonstrated experimentally that configuration 1(gas in the inner compartment of the reactor, where catalytically active phase is) gives better gaseous reactant mass transfer due to the easier access of the gas to the active phase. We propose a complete mathematical model for a threephase catalytic membrane reactor. This model was developed coupling a numerical approach for the mass transfer mechanisms and kinetics determined experimentally, without adjustable parameters. The results obtained by simulation show the versatility of this model: it gives the concentration profiles through the membrane and support as the average concentration profiles for different reactor configurations. The conversion values obtained by simulation are in good agreement with experimental results. Our system shows that the use of catalytic membrane reactors may be an interesting alternative for low-temperature three-phase hydrogenations, giving promising performances. We demonstrated that the mass transfer through the support and the membrane was one of the most important parameters when the separative layer and the support are completely wetted with the liquid. In this case, the relevant parameters are the transmembrane pressure, the membrane thickness, and the mean pore size. We continue our work with chemical reactions requiring a selectivity (regio- or chemioselectivity) induced by the membrane reactor. Acknowledgment This research was supported by the ECC, Brite-Euram Programme, Project No. 0406. Nomenclature CBO= bulk liquid concentration of aniline (m~l-m-~) C%* = equilibriummolar concentration of hydrogen (mol.m-3)
Ci' = dimensionless concentrations C ~ L= dimensionless concentrations in liquid Ci, = molar concentration in the membrane (mol.m-9 Ci, = molar concentration in the support (mo1-m") Dim = effective diffusivity in the membrane (m2.s-1) Di, = effective diffusivity in the support (m2.s-l) H = Henry constant for hydrogen (Pa.m3.mol-1) K = rate constant defined by eq 1 L = length of reactor (m) Mpt = mass of Pt (mg) Pei = Peclet number, Pei = Di/V,L PH = Hydrogen partial pressure (Pa) r = radial coordinate (m) r1 = membrane inner radius (m) r2 = membrane outer radius (m) Ri = consumption rate of ith specimen (mol.rn-%-') Q = liquid flow rate in the loop (m3.s-l) V , = mean velocity of liquid (ms-1) X = axial coordinate (m) z = dimensionless axial coordinate, z = X / L 6 , = membrane thickness (m) 6, = support thickness (m) ym = r1/6, = dimensionless parameter ys = rg/& = dimensionless parameter $H, $B = Thiele moduli, defined by eqs 9 and 10 &, = ( r - r1)/6m= dimensionless radial coordinate in the support (L = (r - ~ ) / ( r-3 r2) = dimensionless radial coordinate in the liquid channel in the configurations 1 and 2 A = rZli-3 = dimensionless parameter from eq 4b
Literature Cited Akyurtlu, J. F.; Akyurtlu, A.; Hamrin, C. E. A Study on the Performance of the Catalytic Porous-wall Three-phase Reactor. Chem. Eng. Commun. 1988,66, 169-187. Cini, P.; Harold, M. P. ExperimentalStudy of the Tubular Multiphase Catalyst. AIChE J. 1991,37 (7), 997-1008. De Vos, R.; Hamrin,C. E. A Cross-Flow Reactor: Theoretical Model for First Order Kinetics. Chem. Eng. Sci. 1982, 37 (12), 17111718. Finlayson, B. A. Nonlinear analysis in Chemical Engineering; McGraw-Hill: New York, 1980; pp 286-291. Harold, M. P.; Cini, P.; Patenaude, B.; Venkatamman, K. The Catalytically Impregnated Ceramic Tube: an Alternative Multiphase Reactor. MChE Symp. Ser. 1989,85,26. Hatziantoniou, V.; Andersson, B. The Segmented Two-Phase Flow Monolithic Catalyst Reactor. An Alternative for Liquid-Phase Hydrogenations. Ind. Eng. Chem. Fundam. 1984,23, 82-88. Mills, P. L.; Dudukovic, M. P. A Comparison of Current Models for Isothermal Trickle-Bed Reactors. In Chemical and Catalytic Reactor Modeling; Dudukovic, M. P., Mills, P. L., Eds.; ACS SymposiumSeries 237;American Chemical Society: Washington, DC, 1984; pp 49-54. Sherwood,T. K.;Pigford,R. P.; Wilke, C. R. Mass Transfer;McGrawHill: New York, 1975; pp 25-28. Torres, M. Etude et ModBlisation d'un RBacteur Membranaire Applique B des RBactions Triphasiques. Ph.D. Dissertation, Universit4 Claude Bernard Lyon I, Lyon, 1993. Uzio, D.; Peureux, J.; Giroir-Fendler,A.; Torres, M.; Dalmon, J. A. Platinum/y AlzOs CatalyticMembranePreparation,Morphological and Catalytic Characterisations.Appl. Catal. 1993,96 (l), 83-98. Received for review February 7, 1994 Revised manuscript received June 1, 1994 Accepted June 22, 1994O
* Abstract published in Advance ACS Abstracts, September 1, 1994.
